Apparatus and process for producing gasoline, olefins and aromatics from oxygenates

ABSTRACT

Apparatuses and processes for converting an oxygenate feedstock, such as methanol and dimethyl ether, in a fluidized bed containing a catalyst to hydrocarbons, such as gasoline boiling components, olefins and aromatics are provided herein.

CROSS-REFERENCE TO RELATED APPLICATIONS

This application claims the benefit of provisional U.S. Ser. Nos.62/062,396 and 62/062,423, both filed Oct. 10, 2014, the entire contentsof each of which are expressly incorporated by reference herein.

FIELD OF THE INVENTION

The present invention relates to converting an oxygenate feedstock, suchas methanol and dimethyl ether, in a fluidized bed containing a catalystto hydrocarbons, such as gasoline boiling components, olefins andaromatics.

BACKGROUND OF THE INVENTION

Processes for converting lower oxygenates such as methanol and dimethylether (DME) to hydrocarbons are known and have become of great interestbecause they offer an attractive way of producing liquid hydrocarbonfuels, especially gasoline, from sources which are not petrochemicalfeeds. In particular, they provide a way by which methanol and DME canbe converted to gasoline boiling components, olefins and aromatics.Olefins and aromatics are valuable chemical products and can serve asfeeds for the production of numerous important chemicals and polymers.Because of the limited supply of competitive petroleum feeds, theopportunities to produce low cost olefins from petroleum feeds arelimited. However, methanol may be readily obtained from coal bygasification to synthesis gas and conversion of the synthesis gas tomethanol by well-established industrial processes. As an alternative,the methanol may be obtained from natural gas or biomass by otherconventional processes.

Available technology to convert methanol and other lower oxygenates tohydrocarbon products utilizes a fixed bed process, such as the processesdescribed in U.S. Pat. Nos. 3,998,899; 3,931,349 and 4,035,430. In thefixed bed process, the methanol is usually first subjected to adehydrating step, using a catalyst such as gamma-alumina, to form anequilibrium mixture of methanol, DME and water. This mixture is thenpassed at elevated temperature and pressure over a catalyst forconversion to the hydrocarbon products which are mainly in the range oflight gas to gasoline. The fixed bed process uses a recycle gas fortemperature control and very large heat transfer to manage low qualityheat, which results in high compression costs and a large heat exchangenetwork. Typically, a fixed bed process is a multi-reactor, unsteadystate operation, which requires a large bore valving system to controlthe process.

In contrast, direct cooling of the reactor in the fluidized bed processeliminates the need for recycle gas for temperature control, whichsimplifies the heat exchange. Further, the fluidized bed process withcontinuous catalyst regeneration is a steady state operation withconstant product yield. Thus, the fluidized bed process requires lowercapital costs and savings on operating expenses compared to the fixedbed process. However, current fluidized bed processes typically have alow product yield. For example, C₅₊ gasoline yield from a fluidized bedprocess ranges from 65 wt % to 75 wt % of hydrocarbons (HC), while theC₅₊ gasoline yield from a fixed bed process ranges from 80 wt % to 90 wt% of HC. Thus, an alkylation unit is usually required to increase C₅₊gasoline yield in a fluidized bed process. Therefore, there is a need toprovide fluidized bed processes for converting oxygenates tohydrocarbons with increased product yields and further, without the useof an alkylation unit.

SUMMARY OF THE INVENTION

It has been found that hydrocarbon product yields can be increasedwithout the need for an alkylation unit by providing apparatuses andprocesses for converting oxygenates in a fluidized reactor bed bystaging the reactor, operating the reactor at a higher pressure andlower temperature, and/or providing a recycle stream, such as lightolefins.

Thus, in one aspect, embodiments of the invention provide a process forconverting an oxygenate feedstock to a C₅₊ gasoline product comprising:feeding the oxygenate feedstock to a fluidized bed reactor underconditions to convert the oxygenate feedstock to a hydrocarbon mixturein a reactor effluent, wherein the fluid bed reactor comprises: (i) acatalyst; and (ii) at least one packing layer; cooling the reactoreffluent comprising the hydrocarbon mixture and condensing a portion ofthe reactor effluent to form a mixed phase effluent; separating themixed phase effluent into an aqueous liquid phase, a hydrocarbon gasphase and a hydrocarbon liquid phase; separating a C⁴⁻ light gascomprising C₂-C₄ olefins and the C₅₊ gasoline product from thehydrocarbon gas phase and the hydrocarbon liquid phase.

In another aspect, embodiments of the invention provide an apparatus forproducing a C₅₊ gasoline product comprising: a fluidized bed reactorcomprising: (i) a fluid inlet for a feedstock; (ii) a catalyst; and(iii) at least one packing layer; a cooler for cooling the reactoreffluent and condensing a portion of the reactor effluent to form amixed phase effluent; a separator for separating the mixed phaseeffluent into a gas hydrocarbon stream, a water stream, and a liquidhydrocarbon stream; a means for transporting the reactor effluent fromthe fluid bed reactor to the separator; at least one fractionatingcolumn for producing the C₅₊ gasoline product; and a means fortransporting the liquid hydrocarbon stream and gas hydrocarbon stream tothe at least one fractionating column.

In still another aspect, embodiments of the invention provide a processfor converting an oxygenate feedstock to olefins comprising: feeding theoxygenate feedstock to a fluidized bed reactor under conditions toconvert the oxygenate feedstock to a hydrocarbon mixture in a reactoreffluent, wherein the fluid bed reactor comprises: (i) a catalyst; and(ii) at least one packing layer; cooling the reactor effluent comprisingthe hydrocarbon mixture and condensing a portion of the reactor effluentto form a mixed phase effluent; separating the mixed phase effluent intoan aqueous liquid phase, a hydrocarbon gas phase and a hydrocarbonliquid phase; separating olefins from the hydrocarbon gas phase and thehydrocarbon liquid phase.

In still another aspect, embodiments of the invention provide a processfor converting an oxygenate feedstock to aromatics comprising: feedingthe oxygenate feedstock to a fluidized bed reactor under conditions toconvert the oxygenate feedstock to a hydrocarbon mixture in a reactoreffluent, wherein the fluid bed reactor comprises: (i) a catalyst; and(ii) at least one packing layer; cooling the reactor effluent comprisingthe hydrocarbon mixture and condensing a portion of the reactor effluentto form a mixed phase effluent; separating the mixed phase effluent intoan aqueous liquid phase, a hydrocarbon gas phase and a hydrocarbonliquid phase; separating aromatics from the hydrocarbon gas phase andthe hydrocarbon liquid phase.

Other embodiments, including particular aspects of the embodimentssummarized above, will be evident from the detailed description thatfollows.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 illustrates a staged fluidized bed methanol to gasoline (MTG)reactor with internal cooling.

FIG. 2 illustrates a staged fluidized bed MTG reactor with externalcooling.

FIG. 3 illustrates an internally cooled fluidized bed MTG process withlight gas recycling.

FIG. 4 illustrates an externally cooled fluidized bed MTG process withlight gas recycling.

FIG. 5 illustrates an internally cooled fluidized bed MTG process withlight gas recycling to a second reactor.

FIG. 6 illustrates an externally cooled fluidized bed MTG process withlight gas recycling to a second reactor.

FIG. 7 illustrates an externally cooled fluidized bed MTG process withlight gas recycling to a catalyst cooler.

FIG. 8 illustrates an internally cooled fluidized bed MTG process withalkylation.

FIG. 9 illustrates an externally cooled fluidized bed MTG process withalkylation.

FIG. 10 illustrates an internally cooled fluidized bed MTG process withlight olefins recycling.

FIG. 11 illustrates an externally cooled fluidized bed MTG process withlight olefins recycling.

FIG. 12 illustrates a staged fluidized bed reactor with internal coolingwith DME feedstock.

FIG. 13 illustrates a staged fluidized bed reactor with external coolingwith DME feedstock.

FIG. 14 illustrates an internally cooled fluidized bed process withlight gas recycling with DME feedstock.

FIG. 15 illustrates an externally cooled fluidized bed process withlight gas recycling with DME feedstock.

FIG. 16 illustrates an internally cooled fluidized bed process withlight gas recycling to a second reactor with DME feedstock.

FIG. 17 illustrates an externally cooled fluidized bed process withlight gas recycling to a second reactor with DME feedstock.

FIG. 18 illustrates an externally cooled fluidized bed process withlight gas recycling to the catalyst cooler with DME feedstock.

DETAILED DESCRIPTION OF THE INVENTION

In various aspects of the invention, apparatuses and processes forconverting oxygenates, such as methanol and DME, in a fluidized bedcomprising a catalyst to hydrocarbons, such as gasoline, olefins andaromatics are provided.

I. Definitions

As used herein, the term “aromatic” refers to unsaturated cyclichydrocarbons have 6 to 18 ring carbon atoms (e.g., 6 to 12 ring carbonatoms), such as but not limited to benzene, toluene, xylenes,mesitylene, ethylbenzenes, cumene, naphthalene, methylnaphthalene,dimethylnaphthalenes, ethylnaphthalenes, acenaphthalene, anthracene,phenanthrene, tetraphene, naphthacene, benzanthracenes, fluoranthrene,pyrene, chrysene, triphenylene, and the like, and combinations thereof.The aromatic may comprise monocyclic, bicyclic, tricyclic, and/orpolycyclic rings (in some embodiments, at least monocyclic rings, onlymonocyclic and bicyclic rings, or only monocyclic rings) and may befused rings.

As used herein, the term “olefin” refers to an unsaturated hydrocarbonchain of 2 to about 12 carbon atoms in length containing at least onecarbon-to-carbon double bond. The olefin may be straight-chain orbranched-chain. Non-limiting examples include ethylene, propylene,butylene, and pentenyl. “Olefin” is intended to embrace all structuralisomeric forms of olefins. As used herein, the term “light olefin”refers to olefins having 2 to 4 carbon atoms (i.e., ethylene, propylene,and butenes).

As used herein, the term “paraffin” refers to a saturated hydrocarbonchain of 1 to about 12 carbon atoms in length, such as, but not limitedto methane, ethane, propane and butane. The paraffin may bestraight-chain or branched-chain. “Paraffin” is intended to embrace allstructural isomeric forms of paraffins. As used herein, the term “lightparaffin” refers to paraffins having 1 to 4 carbon atoms (i.e., methane,ethane, propane and butane).

As used herein, the term “oxygenate” refers to oxygen-containingcompounds having 1 to about 20 carbon atoms, 1 to about 10 carbon atoms,or 1 to about 4 carbon atoms. Exemplary oxygenates include alcohols,ethers, carbonyl compounds, e.g., aldehydes, ketones and carboxylicacids, and mixtures thereof. Particular non-limiting examples ofoxygenates include methanol, ethanol, dimethyl ether, diethyl ether,methylethyl ether, di-isopropyl ether, dimethyl carbonate, dimethylketone, formaldehyde, acetic acid, and the like, and combinationsthereof.

As used herein, the term “C₅₊ gasoline” refers to a compositioncomprising C₅-C₁₂ hydrocarbons and/or having a boiling point rangewithin the specifications for motor gasoline (e.g., from about 100° F.to about 400° F.).

As used herein, the term “coke” refers to a carbonaceous solid or liquidmaterial resulting from conversion of an oxygenate to a hydrocarbon.

As used herein, the term “liquefied petroleum gas” or “LPG” refers to amixture of hydrocarbons in a liquid state, in particular propane andbutane.

II. Converting an Oxygenate to a Hydrocarbon Product

In a first embodiment, an oxygenate feedstock can be fed into afluidized bed reactor comprising a catalyst, and the oxygenate can beconverted into a hydrocarbon product, which can be further separatedinto various hydrocarbon components. The hydrocarbon product yield canbe improved by staging the fluidized bed reactor, by operating thereactor at a relatively high pressure and/or at a relatively lowtemperature, and/or by providing a gas recycle stream to the reactor.

In various aspects, the oxygenated hydrocarbon feedstock can comprisemethanol, DME, or a mixture thereof. The methanol can be obtained fromcoal, natural gas and biomass by conventional processes.

In various aspects, the hydrocarbon product can comprise C₅₊ gasoline,aromatics, and/or olefins.

In one aspect, the gas recycle stream can comprise olefins.

III. Structured Packing

In any embodiment, the fluidized bed reactor can include at least onelayer of structured packing as a staging baffle. A deep fluidized beddesign can be used for the reactor due to the weight hourly spacevelocity (WHSV) required by the chemical reactions. However, a deepfluidized bed can be prone to gas back-mixing and gas by-pass.Therefore, it can be important to minimize gas back-mixing and gasby-pass to maintain oxygenate conversion and maximize product yield. Thegas back-mixing and gas by-pass can be minimized by installing at leastone layer of structured packing which functions as a staging baffle inthe fluidized bed reactor. Advantageously, the fluidized bed reactor caninclude at least two layers of structured packing FIGS. 1, 2, 12, and 13show a fluid bed reactor with two layers of structured packing as twostructured baffles. However, in various aspects, the fluidized bedreactor can include from one to eight layers of structured packing.

An example of the structured packing is a one foot thick layer ofKoch-Glitsch KFBE IIB, which separates the dense fluid bed into multiplestages. Structured packing is commonly used in distillation towers inseparation processes. This type of packing can be useful because of itshigh open area for both gas and catalyst solids to pass through and itscapability to control bubble sizes. When larger bubbles from a lowerstage reach the staging baffles, gas can be redistributed by thestructured packing and form smaller bubbles into the next higher stage.

IV. Production of Gasoline

In a methanol to gasoline (MTG) process, methanol can first bedehydrated to form dimethyl ether. The methanol and/or dimethyl ethercan then be converted in a series of reactions that result in formationof a hydrocarbon mixture that can comprise aromatics, paraffins, andolefins, among other types of hydrocarbon products. This mixture may beseparated into a LPG fraction and a high-quality gasoline fraction,e.g., comprising aromatics, paraffins, and olefins.

In one embodiment, the oxygenate feedstock comprises methanol, which isfed into a fluidized bed reactor and converted to gasoline boilingcomponents in an MTG process. In some embodiments, the methanol can beobtained from coal with a water content up to 15 wt %, for example from5 wt % to 10 wt %, and/or from natural gas with a water content up to 40wt %, for example from 30 wt % to 40 wt %.

Traditionally, C₅₊ gasoline yield in MTG processes can be in the rangeof 65-72 wt %, based on the feed, when relatively high temperatures(˜715-800° F.) and/or relatively low pressures (˜25-45 psig) are used inthe MTG reactor. However, it is believed that the C₅₊ gasoline yield ofthe fluid bed MTG process can be improved to at least about 75 wt %, forexample at least about 80 wt %, about 80-90 wt %, about 80-85 wt %,about 85-90 wt %, or about 86-95 wt %, as compared to the feed,advantageously without the need for an alkylation unit by staging thereactor, by operating the reactor at a higher pressure and lowertemperature, and/or by recycling the light olefins. The spent catalystsfrom the reactor can be transferred to the regenerator to regenerate thecatalyst by burning the coke off. The regenerated catalysts can then betransferred back to the reactor.

Alternatively or additionally, the oxygenate feedstock can comprise DME,which can be fed into a fluidized bed reactor and converted to gasolineboiling components. The DME to gasoline process can achieve a C₅₊gasoline yield of greater than 70 wt %, for example at least about 75 wt%, at least about 80 wt %, about 75-95 wt %, about 75-90 wt %, about80-90 wt %, about 80-85 wt %, about 85-90 wt %, or about 86-95 wt %, ascompared to the feed, advantageously without the need for an alkylationunit by staging the reactor, by operating the reactor at a higherpressure and lower temperature, and/or by recycling the light olefins.The spent catalysts from the reactor can be transferred to theregenerator to regenerate the catalyst by burning the coke off. Theregenerated catalysts can then be transferred back to the reactor.

In any embodiment, the fluidized bed reactor can include at least onelayer of structured packing as a staging baffle. In various aspects, thefluid bed reactor can include from one to eight layers of structuredpacking Advantageously, the fluid bed reactor can include at least twolayers of structured packing By including structured packing, it isbelieved that the C₅₊ gasoline yield can be further improved by at least2-4 wt % for the MTG process.

A. Cooling the Reactor

The conversion of methanol and/or DME to gasoline boiling components isa highly exothermic reaction. For example, the MTG process releasesapproximately 750 BTU of heat per pound of methanol. Thus, it can oftenbe necessary to cool the fluidized bed reactor.

In one embodiment, the fluidized bed reactor can be internally cooled,such as shown in FIGS. 1, 3, 5, 8, 10, 12, 14, and 16. For example, aheat exchanger can be present in one or more stages. FIGS. 1, 3, 5, 8,10, 12, 14, and 16 show heat exchangers in each stage. The internal heatexchangers can function not only to remove the heat from the reactor butalso as internal baffles, operating to break up large bubbles and thusreduce gas by-pass. With internal heat exchangers, controlling thetemperature at each stage is also achievable, which can provide theability to adjust the process operation to improve (maximize) desiredproduct (e.g., C₅₊ gasoline) yield.

Additionally or alternatively, the fluidized bed reactor can beexternally cooled, as shown in FIGS. 2, 4, 6, 7, 9, 11, 13, 15, 17, and18. For example, a catalyst cooler can be installed for removing theheat from the reactor by circulating the catalyst between the reactorand the cooler, as shown in FIGS. 2, 4, 6, 7, 9, 11, 13, 15, 17, and 18.

With in-bed heat exchangers and/or external catalyst cooler(s), arelatively uniform temperature distribution within the operating rangecan be achieved in the fluidized bed process. While the internal coolingoption can be easier to operate, the external cooling option can providemore flexibility for operation and a less complicated construction,especially for a large scale unit.

B. Operating Conditions

The fluidized bed reactor can be operated at pressure from about 25 psigto about 400 psig, for example from about 75 psig to about 400 psig,from about 75 psig to about 300 psig, from about 75 psig to about 200psig, from about 100 psig to about 400 psig, from about 100 psig toabout 300 psig, from about 100 psig to about 200 psig, from about 150psig to about 350 psig, at about 150 psig, at about 200 psig, or atabout 250 psig. The fluidized bed reactor can be operated at atemperature from about 500° F. to about 900° F., for example from about550° F. to about 900° F., from about 600° F. to about 900° F., fromabout 700° F. to about 900° F., from about 500° F. to about 750° F.,from about 500° F. to about 700° F., from about 500° F. to about 650°F., from about 600° F. to about 700° F., from about 550° F. to about700° F., from about 550° F. to about 650° F., from about 550° F. toabout 600° F., at about 550° F., at about 600° F., at about 650° F., orat about 700° F. Further, the methanol WHSV can be from about 0.2kg/kg-hr to about 3.0 kg/kg-hr, for example from about 0.5 kg/kg-hr toabout 2.5 kg/kg-hr, from about 1 kg/kg-hr to about 2.0 kg/kg-hr, atabout 1.7 kg/kg-hr, or at about 1.5 kg/kg-hr, during operation. Underthese operating conditions, it is believed that the desired product(e.g., C₅₊ gasoline) yield can be improved by at least 4-6 wt % (for theMTG process).

C. Catalysts

The conversion reactions described herein typically utilize a catalyst.Useful catalyst compositions for MTG processes can comprise boundzeolite catalysts and unbound zeolite catalysts.

Generally, the zeolite employed in the present catalyst composition cantypically have a silica to alumina molar ratio of at least 40, e.g.,from about 40 to about 200. Additionally or alternately, the zeolite cancomprise at least one medium pore aluminosilicate zeolite having aConstraint Index of 1-12 (as defined in U.S. Pat. No. 4,016,218).Suitable zeolites can include, but are not necessarily limited to,ZSM-5, ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-35, ZSM-48, and the like, aswell as combinations thereof. ZSM-5 is described in detail in U.S. Pat.Nos. 3,702,886 and RE 29,948. ZSM-11 is described in detail in U.S. Pat.No. 3,709,979. ZSM-12 is described in U.S. Pat. No. 3,832,449. ZSM-22 isdescribed in U.S. Pat. No. 4,556,477. ZSM-23 is described in U.S. Pat.No. 4,076,842. ZSM-35 is described in U.S. Pat. No. 4,016,245. ZSM-48 ismore particularly described in U.S. Pat. No. 4,234,231. In certainembodiments, the zeolite can comprise, consist essentially of, or beZSM-5. The ZSM-5 can have a silica to alumina ratio of 55:1.

When used in the present catalyst composition, the zeolite canadvantageously be present at least partly in the hydrogen form.Depending on the conditions used to synthesize the zeolite, this mayimplicate converting the zeolite from, for example, the alkali (e.g.,sodium) form. This can readily be achieved, e.g., by ion exchange toconvert the zeolite to the ammonium form, followed by calcination in airor an inert atmosphere at a temperature from about 400° C. to about 700°C. to convert the ammonium form to the active hydrogen form. If anorganic structure directing agent is used in the synthesis of thezeolite, additional calcination may be desirable to remove the organicstructure directing agent.

The catalysts described herein can be pretreated with steam prior to usein the reactor.

To enhance the steam stability of the zeolite without excessive loss ofits initial acid activity, the present catalyst composition can containphosphorus in an amount between about 0.01 wt % and about 3 wt %elemental phosphorus, e.g., between about 0.05 wt % and about 2 wt %, ofthe total catalyst composition. The phosphorus can be added to thecatalyst composition at any stage during synthesis of the zeolite and/orformulation of the zeolite and binder into the catalyst composition.Generally, phosphorus addition can be achieved by spraying and/orimpregnating the final catalyst composition (and/or a precursor thereto)with a solution of a phosphorus compound. Suitable phosphorus compoundscan include, but are not limited to, phosphinic [H₂PO(OH)], phosphonic[HPO(OH)₂], phosphinous, phosphorus, and phosphoric [PO(OH)₃] acids,salts and esters of such acids, phosphorus halides, and the like, andcombinations thereof. After phosphorus treatment, the catalyst cangenerally be calcined, e.g., in air at a temperature from about 400° C.to about 700° C. to convert the phosphorus to an oxide form.

In one embodiment, the catalyst is modified with up to 3 wt %phosphorous for improved stability.

Additionally or alternatively, the catalyst composition can include upto 80% clay by weight, for example up to 50 wt % clay, up to 40 wt %clay, or up to 30 wt % clay.

D. Recycling Light Olefins

In some embodiments, the desired product (e.g., C₅₊ gasoline) yield canbe improved by recycling C⁴⁻ light gas (e.g., which light gas cancomprise olefins to convert ethylene, propylene, and butenes to C₅₊gasoline). For a fixed bed MTG process, the C₅₊ gasoline yield can be inthe range of about 80-90 wt % with a light gas recycle ratio of about6-9. For a fluid bed MTG process, by recycling the C⁴⁻ light gas with arecycle ratio of up to 3, it is believed that the C₅₊ gasoline yield canbe further improved by at least about 8-12 wt %. Additionally oralternately, recycling light gas can improve the stabilization of thefluidized bed reactor hydrodynamics. The olefins can be recycled to themain reactor (FIGS. 3-4 and 14-15) or to a second reactor (FIGS. 5-7 and16-18).

Table 1 can represent a product yield for a fluid bed MTG process. Asshown in Table 1, C₅₊ gasoline yield is about 67 wt %, and light olefinyield (including ethylene, propylene, and butenes) is about 17 wt %. Byrecycling the light olefins and converting all of them to C₅₊ gasoline,the potential C₅₊ gasoline yield can be about 84 wt %, which can thencompare more favorably to fixed bed MTG gasoline yields. Considering thesteady state operation, uniform temperature distribution in the reactorwith direct cooling and smaller recycle heater and compressor, the fluidbed MTG process can provide both the capital and operating costadvantages over the fixed bed MTG process.

TABLE 1 Typical Product Yield of a Fluid Bed MTG Process Fluid Bed MTGYield Hydrocarbon Product, wt % of HC Light Gas 2.7 Ethylene 5.4 Propane3.5 Propylene 5.4 i-Butane 8.5 n-Butane 1.5 Butenes 5.8 C5+ Gasoline67.2 Potential C5+ Gasoline w Alkylate 91.2 Potential C5+ Gasoline w/oAlkylate 83.8

In one embodiment, C⁴⁻ light gas comprising olefins can be recycled tothe main reactor to convert ethylene, propylene, and butenes to C₅₊gasoline, as shown in FIGS. 3 and 4. Heated methanol feed can be fed tothe bottom of the reactor. Reactor vapor can be separated from catalystby a set of two stage cyclones. Reactor effluent can be sent to finescollection equipment, e.g., KBR CycloFines™, to remove catalyst fines.The reactor effluent can be further cooled and partially condensedagainst incoming methanol feed, and then the mixed phase effluent can besent to a water separator where the condensed aqueous phase can beseparated and sent to wastewater treatment. Separator vapor andhydrocarbon liquid can be sent to a stabilizer, where C⁴⁻ light gas canbe separated from C₅₊ product. The C⁴⁻ light gas can be recycled back tothe reactor where the light olefins can be converted to C₅₊ gasoline.The C₅₊ gasoline product can be used immediately or sent to storage forlater use. The spent catalysts from the reactor can be transferred to aregenerator to regenerate the catalyst, e.g., by burning the coke off.The regenerated catalysts can then be transferred back to the reactor.In the alternative, a DME feed can be fed to the reactor instead of orin addition to methanol feed, as shown in FIGS. 14 and 15.

Additionally or alternatively, after the mixed phase effluent is sent tothe water separator where the condensed aqueous phase can be separatedand sent to wastewater treatment, the separator vapor and hydrocarbonliquid can be sent to one or multiple dividing wall columns wheremultiple (e.g., seven) streams (e.g., including light gas, C₂,propylene, propane, butenes, butanes, and C₅₊ product) can be divided,as shown in FIG. 10. In MTG processes, C₂, propylene, and butenes can becombined and advantageously recycled back to the reactor to be furtherconverted to C⁵⁻ gasoline. Propane and butanes can be combined as LPG.C₅₊ gasoline product can be used immediately or sent to storage forlater use.

Additionally or alternatively to recycling to the main reactor, C⁴⁻light gas comprising olefins can be recycled to a different (second)reactor, e.g., to convert ethylene, propylene, and butanes to C₅₊gasoline, as shown in FIGS. 5 and 6. Heated methanol feed can be fed tothe bottom of the reactor. Reactor vapor can be separated from catalystby a set of two stage cyclones. Reactor effluent can be sent to finescollection equipment, e.g., KBR CycloFines™, to remove catalyst fines.The reactor effluent can be further cooled and partially condensedagainst incoming methanol feed, and then the mixed phase effluent can besent to a water separator where the condensed aqueous phase can beseparated and sent to wastewater treatment. Separator vapor andhydrocarbon liquid can be sent to a stabilizer where C⁴⁻ light gascomprising olefins can be separated from C⁵⁻ product. The light C⁴⁻gases can be recycled to a different (second) reactor, e.g., whereethylene, propylene, and butenes can be converted to C₅₊ gasoline. Theheat in the second reactor from the exothermic reaction can be removedby internal heat exchangers. The effluent from the second reactor afterprocessing through a set of two stage cyclones can be sent to finescollection equipment, e.g., KBR Cyclofines™, to remove catalyst fines.The effluent from the second reactor can be further cooled and sent to asecond water separator where condensed aqueous phase is separated andsent to wastewater treatment. Separator gas and hydrocarbon liquid fromthe second separator can be combined with C₅₊ product from thestabilizer and sent to a de-ethanizer fractionating column, where C²⁻light gas can be separated from C₃₊ product. The C³⁻ product can be sentto a de-butanizer fractionating column where the LPG can be separatedfrom C₅₊ gasoline product. The C₅₊ gasoline product can be usedimmediately or sent to storage for later use. The spent catalysts fromthe reactor can be transferred to a regenerator to regenerate thecatalyst, e.g., by burning the coke off. The regenerated catalysts canthen be transferred back to the reactor. In certain embodiments, a DMEfeed can be fed to the reactor instead of or in addition to methanolfeed, as shown in FIGS. 16 and 17.

In another embodiment, C⁴⁻ light gas comprising olefins can be recycledto the catalyst cooler to convert ethylene, propylene, and butanes toC₅₊ gasoline, as shown in FIG. 7. The heated methanol feed can be fed tothe bottom of the reactor. Heat from the exothermic reaction can beremoved by the external catalyst cooler. Reactor vapor can be separatedfrom catalyst by a set of two stage cyclones. Reactor effluent can besent to fines collection equipment, e.g., KBR CycloFines™, to removecatalyst fines. The reactor effluent can further be cooled and partiallycondensed against the incoming methanol feed, and then the mixed phaseeffluent can be sent to water separator where the condensed aqueousphase can be separated and sent to wastewater treatment. Separator vaporand hydrocarbon liquid can be sent to a stabilizer where C⁴⁻ gases areseparated from C₅₊ product. C⁴⁻ gases can be recycled to a catalystcooler where ethylene, propylene, and butenes are converted to C₅₊gasoline. The effluent from the catalyst cooler after processing througha set of two stage cyclones can be sent to fines collection equipment,e.g., KBR Cyclofines™, to remove catalyst fines. The effluent from thecatalyst cooler can further be cooled and sent to a second waterseparator where condensed aqueous phase can be separated and sent towastewater treatment. Separator gas and hydrocarbon liquid from thesecond separator can be combined with C₅₊ product from stabilizer andsent to a de-ethanizer fractionating column where C²⁻ light gas can beseparated from C₃₊ product. The C₃₊ product can be sent to ade-butanizer fractionating column where the LPG can be separated fromC₅₊ gasoline product. C₅₊ gasoline product can be used immediately orsent to storage for later use. The spent catalysts from the reactor canbe transferred to a regenerator to regenerate the catalyst by burningthe coke off. The regenerated catalysts can then be transferred back tothe reactor. In certain embodiments, a DME feed can be fed to thereactor instead of or in addition to methanol feed, as shown in FIG. 18.

As discussed previously, heat from the exothermic reaction in thereactor can be removed internally, for example, by internal heatexchangers, as shown in FIGS. 3, 5, 10, 14, and 16.

Additionally or alternatively, heat from the exothermic reaction in thereactor can be removed externally, for example, by an external catalystcooler, as shown in FIGS. 4, 6, 7, 9, 11, 15, 17, and 18.

E. Alkylation

To further improve the C₅₊ gasoline yield, an alkylation unit canoptionally be included in the fluidized bed process to convert isobutaneto C₅₊ gasoline. As shown above in Table 1, C₅₊ gasoline yield is about67 wt %. With an optional alkylation unit, by recycling the isobutane,as well as the alkenes, and converting all of them to C₅₊ gasoline, theC₅₊ gasoline yield of a fluidized bed MTG process could be beyond 90 wt%.

In a further embodiment, an alkylation unit can be included to convertisobutene, propylene, and butenes to C₅₊ gasoline, as shown in FIGS. 8and 9. The heated methanol feed can be fed to the bottom of the reactor.Reactor vapor can be separated from the catalyst by a set of two stagecyclones. Reactor effluent can be sent to fines collection equipment,e.g., KBR CycloFines™, to remove catalyst fines. The reactor effluentcan be further cooled and partially condensed against the incomingmethanol feed, and then the mixed phase effluent can be sent to a waterseparator where the condensed aqueous phase can be separated and sent towastewater treatment. Separator vapor and hydrocarbon liquid can be sentto a de-ethanizer fractionating column where C²⁻ light gas can beseparated from C₃₊ product. The C₃₊ product can be sent to ade-butanizer fractionating column where the C₃/C₄ gases can be separatedfrom C₅₊ gasoline product. C₅₊ gasoline product can be used immediatelyor sent to storage for later use. C₃/C₄ gases can be sent to analkylation unit to convert isobutene, propylene, and butenes to C₅₊gasoline.

V. Production of Olefins

In another embodiment, an oxygenate is fed into a fluidized bed reactorand converted to olefins. In various aspects, the oxygenate is methanol,DME or a mixture thereof. The olefin yield of the fluid bed process canbe improved by staging the reactor, operating the reactor at a higherpressure and lower temperature, and/or by providing a recycle gasstream. The spent catalysts from the reactor are transferred to theregenerator to regenerate the catalyst by burning the coke off. Theregenerated catalysts are then transferred back to the reactor.

In any embodiment, the fluidized bed reactor can include at least onelayer of structured packing as a staging baffle. In various aspects, thefluid bed reactor can include from one to eight layers of structuredpacking Advantageously, the fluid bed reactor can include at least twolayers of structured packing

In one embodiment, the fluidized bed reactor can be internally cooled,for example with a heat exchanger that is present in at least one oreach stage for cooling. Additionally or alternatively, the fluidized bedreactor is externally cooled, for example, with a catalyst coolerinstalled for removing the heat from the reactor by circulating thecatalyst between the reactor and the cooler.

In various aspects, olefins are produced according to the processesdescribed above for converting methanol and/or DME into C₅₊ gasoline.Heated methanol feed and/or DME feed can be fed to the bottom of thereactor. Reactor vapor is separated from catalyst by a set of two stagecyclones. Reactor effluent is sent to fines collection equipment, e.g.,KBR CycloFines™, to remove catalyst fines. The reactor effluent isfurther cooled and partially condensed against incoming methanol and/orDME feed and then the mixed phase effluent is sent to a water separatorwhere the condensed aqueous phase is separated and sent to wastewatertreatment. Separator vapor and hydrocarbon liquid are sent to astabilizer, where C⁴⁻ light gas comprising olefins is separated from C₅₊product. The C₅₊ gasoline product is sent to storage. The spentcatalysts from the reactor are transferred to a regenerator toregenerate the catalyst by burning the coke off. The regeneratedcatalysts are then transferred back to the reactor.

Additionally or alternatively, a recycle gas stream can be sent toeither the reactor, a second reactor, or the catalyst cooler forincreasing the yield of olefins.

A. Operating Conditions

The fluidized bed reactor can be operated at a pressure from about 3psig to about 450 psig, for example from about 75 psig to about 400psig, from about 75 psig to about 300 psig, from about 75 psig to about200 psig, from about 100 psig to about 400 psig, from about 100 psig toabout 300 psig, from about 100 psig to about 200 psig, from about 150psig to about 350 psig, at about 150 psig, at about 200 psig, or atabout 250 psig. The fluidized bed reactor can be operated at atemperature from about 500° F. to about 1100° F., for example from about700° F. to about 1100° F., from about 800° F. to about 1100° F., fromabout 900° F. to about 1100° F., from about 650° F. to about 1050° F.,from about 650° F. to about 1000° F., from about 750° F. to about 1050°F., from about 800° F. to about 1050° F., from about 850° F. to about1000° F., from about 850° F. to about 1050° F., from about 950° F. toabout 1100° F., at about 950° F., at about 1000° F., or at about 1050°F. Further, the WHSV can be from about 0.1 kg/kg-hr to about 200kg/kg-hr, for example from about 0.5 kg/kg-hr to about 25 kg/kg-hr, fromabout 1 kg/kg-hr to about 20 kg/kg-hr, or at about 1.6 kg/kg-hr, duringoperation.

B. Catalysts

A catalyst is used in the process described herein, which is useful forthe conversion of oxygenate feeds to olefins.

The catalysts described herein can be pretreated with steam prior to usein the reactor and may contain up to 3 wt % of an element to conveysteam stability, such as phosphorus and/or zinc.

In various aspects, a catalyst composition comprising a class ofzeolites described in detail in U.S. Pat. Nos. 4,025,575 and 4,083,889,incorporated herein by reference, are useful for conversion of oxygenatefeeds to olefins. The class of zeolites has a silica to alumina ratio ofat least about 12, at least about 40, or at least about 70, and astructure providing constrained access to the crystalline free space.Additionally or alternatively, the class of zeolites has a crystalframework density, in the dry hydrogen form of not substantially belowabout 1.6 grams per cubic centimeter. Additionally or alternatively, theclass of zeolites has a constraint index from about 1 to about 12.Examples of suitable zeolites include, but are not limited to ZSM-5,ZSM-11, ZSM-12, ZSM-21, ZSM-35, ZSM-38 and other similar material.

The zeolites useful as catalysts may be in the hydrogen form or they maybe base exchanged or impregnated to contain ammonium or a metal cationcomplement. It is desirable to calcine the zeolite after base exchange.The metal cations that may be present include any of the cations of themetals of Groups I through VIII of the periodic table. However, in thecase of Group IA metals, the cation content should in no case be solarge as to substantially eliminate the activity of the zeolite for thecatalysis being employed.

Additionally or alternatively, the catalyst composition can be used inthe presence of 2 moles to 20 moles, for example 3 moles to 10 moles, ofsteam per mol of methanol feed, as described in U.S. Pat. No. 4,083,889.The steam diluent may be provided directly by injecting the requisiteamount of water or steam into the reaction zone; or it may be providedtotally or in part by water mixed with the methanol feed, it beingunderstood that the water forms steam in the reaction zone at theprescribed reaction conditions. Further, the steam diluent may besupplemented with an inert diluent selected from the group consisting ofhydrogen, helium, nitrogen, carbon dioxide, a C₁ to C₇ hydrocarbon andflue gas. In such a case, up to 20 total moles of steam plus inertdiluent may be used.

In other aspects, zeolites of the erionite-offretite family as describedin detail in U.S. Pat. No. 4,079,095, incorporated herein by reference,are useful for conversion of oxygenate feeds to olefins. Included withinthis group of zeolites is erionite, both synthetic and natural,offretite, both synthetic and natural, zeolite T and zeolite ZSM-34.Additionally or alternatively, these zeolites may be compounded with aporous matrix material, such as alumina, silica-alumina,silica-magnesia, silica-zirconia, silica-thoria, silica-beryllia,silica-titania, as well as ternary combinations, such assilica-alumina-thoria, silica-alumina-zirconia, silica-alumina-magnesiaand silica-magnesia-zirconia. The matrix may be in the form of a cogel.The relative proportions of finely divided zeolite and inorganic oxidegel matrix may vary widely with the zeolite content ranging from betweenabout 1 to about 99 percent by weight and more usually in the range ofabout 5 to about 80 percent by weight of the composite.

In another aspect, catalyst compositions comprising the zeolite, ZSM-48,as described in detail in U.S. Pat. No. 4,476,338, incorporated hereinby reference, are useful for conversion of oxygenate feeds to olefins.

In still other aspects, catalyst compositions comprisingsilicoaluminophosphate (SAPO) molecular sieves as described in detail inU.S. Pat. No. 4,677,242, incorporated herein by reference, are usefulfor conversion of oxygenate feeds to olefins. Examples of useful SAPOmolecular sieves include but are not limited to SAPO-5, SAPO-11,SAPO-16, SAPO-17, SAPO-20, SAPO-31, SAPO-34, SAPO-35, SAPO-37, SAPO-40,SAPO-41, SAPO-42 and SAPO-44.

In still other aspects, catalyst compositions comprising non-zeoliticmolecular sieves as described in detail in U.S. Pat. No. 4,752,651,incorporated herein by reference, are useful for conversion of oxygenatefeeds to olefins. Examples of useful non-zeolitic molecular sievesinclude but are not limited to ELAPSO, metal aluminophosphates (MeAPOswhere “Me” is at least one of Mg, Mn, Co and Zn), ferroaluminophosphates(FeAPO or FAPO), titanium aluminophosphates (TAPO), ELAPO, TiAPSO,MgAPSO, MnAPSO, CoAPSO, ZnAPSO, FeAPSO, CoMnAPSO and CoMnMgAPSOmolecular sieves as described in U.S. Pat. No. 4,752,641.

VI. Production of Aromatics

In certain embodiments, an oxygenate is fed into a fluidized bed reactorand converted to aromatics. In various aspects, the oxygenate cancomprise or be methanol and/or DME. The aromatic yield of the fluid bedprocess can be improved by staging the reactor, operating the reactor ata higher pressure and lower temperature, and/or by providing a recyclegas stream. The spent catalysts from the reactor can be transferred tothe regenerator to regenerate the catalyst by burning the coke off. Theregenerated catalysts can then be transferred back to the reactor.

In any embodiment, the fluidized bed reactor can include at least onelayer of structured packing as a staging baffle. In various aspects, thefluid bed reactor can include from one to eight layers of structuredpacking Advantageously, the fluid bed reactor can include at least twolayers of structured packing.

In certain embodiments, the fluidized bed reactor can be internallycooled, for example with a heat exchanger that is present in at leastone or each stage for cooling. Additionally or alternatively, thefluidized bed reactor can be externally cooled, for example, with acatalyst cooler installed for removing the heat from the reactor bycirculating the catalyst between the reactor and the cooler.

In various aspects, aromatics can be produced according to the processesdescribed above for converting methanol and DME into C₅₊ gasoline.Heated methanol feed and/or DME feed can be fed to the bottom of thereactor. Reactor vapor can be separated from catalyst by a set of twostage cyclones. Reactor effluent can be sent to fines collectionequipment, e.g., KBR CycloFines™, to remove catalyst fines. The reactoreffluent can be further cooled and partially condensed against incomingmethanol and/or DME feed, and then the mixed phase effluent can be sentto a water separator where the condensed aqueous phase can be separatedand optionally sent to wastewater treatment. Aromatics can be separatedfrom the separator vapor and hydrocarbon liquid and sent to storage. Thespent catalysts from the reactor can be transferred to a regenerator toregenerate the catalyst by burning the coke off. The regeneratedcatalysts can then be transferred back to the reactor.

Additionally or alternatively, a recycle gas stream can be sent to thereactor, to a second reactor, or to the catalyst cooler for increasingthe yield of aromatics.

A. Operating Conditions

The fluidized bed reactor can be operated at pressure from about 3 psigto about 450 psig, such as about 35 psig. The fluidized bed reactor canbe operated at a temperature of about 500° F. to about 1100° F., such asabout 1000° F. Further, the WHSV can be from about 0.1 to about 200, forexample about 1.6, during operation.

B. Catalysts

A zeolite catalyst composition can be used for the conversion ofoxygenate feeds to aromatics. While some catalyst compositions include abinder, in other cases, the catalyst composition may be referred to asbeing “self-bound” or “unbound.” The terms “unbound” and “self-bound”are intended to be synonymous and mean that such a catalyst compositionis free of any inorganic oxide binders, such as alumina and/or silica,which are frequently combined with zeolite catalysts to enhance theirphysical properties.

The catalysts described herein can be pretreated with steam prior to usein the reactor.

A zeolite employed in the present catalyst composition can generallycomprise at least one medium pore aluminosilicate zeolite or silicaaluminophosphate (SAPO) having a Constraint Index of 1-12. TheConstraint Index may be ≤about 12, ≤ about 11, ≤about 10, ≤about 9,≤about 8, ≤about 7, ≤about 6, ≤about 5, ≤about 4, ≤ about 3, or ≤about2. Additionally or alternatively, the Constraint Index may be about≥about 11, ≥about 10, ≥about 9, ≥about 8, ≥about 7, ≥about 6, ≥about 5,≥about 4, ≥about 3, ≥about 2, or ≥about 1. In any embodiment, theConstraint Index may be 1 to about 10, 1 to about 8, 1 to about 6, 1 toabout 5, 1 to about 3, about 2 to about 11, about 3 to about 10, about 4to about 9, or about 6 to about 9, etc. Constraint Index is determinedas described in U.S. Pat. No. 4,016,218, incorporated herein byreference for details of the method. Suitable zeolites include zeoliteshaving an MFI or MEL framework, such as ZSM-5 or ZSM-11.

Some useful catalysts compositions can include a zeolite having astructure wherein there is at least one 10-member ring channel and nochannel of rings having more than 10 members. Some such molecular sievesmay be referred to as having a framework type or topology of EUO, FER,IMF, LAU, MEL, MRI, MFS, MTT, MWW, NES, PON, SFG, STF, STI, TUN, or PUN.Particularly useful zeolites can have a BEA, MFI, or MEL framework type.

Non-limiting examples of SAPOs useful herein include one or acombination of SAPO-5, SAPO-8, SAPO-11, SAPO-16, SAPO-17, SAPO-18,SAPO-20, SAPO-31, SAPO-34, SAPO-35, SAPO-36, SAPO-37, SAPO-40, SAPO-41,SAPO-42, SAPO-44, SAPO-47, and SAPO-56.

Particular other zeolites useful in embodiments of the invention caninclude ZSM-5, ZSM-11; ZSM-12; ZSM-22; ZSM-23; ZSM-34, ZSM-35; ZSM-48;ZSM-57; and/or ZSM-58. Other useful zeolites may additionally oralternately include MCM-22, PSH-3, SSZ-25, MCM-36, MCM-49 or MCM-56,with MCM-22. In any embodiment the zeolite may comprise or be ZSM-5 orZSM-11. ZSM-5 is described in detail in U.S. Pat. Nos. 3,702,886 and RE29,948. ZSM-11 is described in detail in U.S. Pat. No. 3,709,979. ZSM-5can be particularly useful.

Generally, a zeolite having the desired activity can have a silicon toaluminum molar ratio of about 10 to about 300. In any embodiment, themolar ratio of silicon to aluminum may be ≤about 300, ≤about 200, ≤about150, ≤about 125, ≤ about 100, ≤about 80, ≤about 60, ≤about 50, ≤about40, ≤ about 30, ≤about 25, ≤ about 20, ≤about ≤15 or about ≤10.Additionally or alternatively, the molar ratio of silicon to aluminummay be ≥about 10, ≥about 15, ≥about 20, ≥about 25, ≥about 30, ≥about 40,≥about 50, ≥about 60, ≥about 80, ≥about 100, ≥about 125, ≥about 150, or≥about 200; e.g., 20 to about 200, about 30 to about 100, about 40 toabout 80, about 50 to about 50, about 15 to about 100, or about 20 toabout 40.

In some preferred aspects, the silicon to aluminum ratio can be at leastabout 20, such as at least about 30 or at least about 40. In suchembodiments, the silicon to aluminum ratio can optionally be about 80 orless, such as about 60 or less, or about 50 or less, or about 40 orless. Typically, reducing the molar ratio of silicon to aluminum in azeolite can result in a zeolite with a higher acidity, and therefore inhigher activity for cracking of hydrocarbon or hydrocarbonaceous feeds,such as petroleum feeds. However, with respect to conversion ofoxygenates to aromatics, such increased cracking activity may not bebeneficial, and instead may result in increased formation of residualcarbon or coke during the conversion reaction. Such residual carbon candeposit on the zeolite catalyst, leading to deactivation of the catalystover time. Having a molar ratio of silicon to aluminum of ≥about 40,such as ≥about 50 or ≥about 60, can reduce and/or minimize the amount ofadditional residual carbon formed due to the acidic or cracking activityof the catalyst.

It is noted that the molar ratio described herein is a ratio of siliconto aluminum. If a corresponding ratio of silica to alumina is described,the corresponding ratio of silica (SiO₂) to alumina (Al₂O₃) would betwice as large, due to the presence of two aluminum atoms in eachalumina stoichiometric unit. Thus, a molar ratio of silicon to aluminumof 10 corresponds to a silica to alumina ratio of 20.

When used in the present catalyst compositions, the zeolite can bepresent at least partly in the hydrogen (acid, active) form. Dependingon the conditions used to synthesize the zeolite, this may correspond toconverting the zeolite from, for example, the sodium form. This canreadily be achieved, for example, by ion exchange to convert the zeoliteto the ammonium form followed by calcination in air or an inertatmosphere at a temperature of about 400° C. to about 700° C. to convertthe ammonium form to the (active) hydrogen form.

Zeolite catalyst compositions can include and/or be enhanced by atransition metal. Catalyst compositions herein can include a Group 10-12element, or combinations thereof, of the Periodic Table. Exemplary Group10 elements can include, nickel, palladium, and/or platinum,particularly nickel. Exemplary Group 11 elements can include copper,silver, and/or gold, particularly copper. Exemplary Group 12 elementscan include, e.g., zinc and/or cadmium. Advantageously, the transitionmetal can comprise or be a Group 12 metal from the periodic table(sometimes designated as Group IIB) such as Zn and/or Cd. In particularembodiments, nickel, copper and/or zinc, particularly zinc, may be used.The Group 10-12 element can be incorporated into the zeolite by anyconvenient method, such as by impregnation or by ion exchange. Aftersuch incorporation, the Group 10-12 element-enhanced catalyst can betreated in an oxidizing environment (air) or an inert atmosphere at atemperature of about 400° C. to about 700° C.

The amount of Group 10-12 element can be related to the molar amount ofaluminum present in the zeolite. In certain embodiments, the molar ratioof the Group 10-12 element to aluminum in the zeolite can be about 0.1to about 1.3. For example, the molar ratio of the Group 10-12 element toaluminum in the zeolite can be about ≥0.1, e.g., ≥about 0.2, ≥about 0.3,or ≥about 0.4. Additionally or alternately, the molar ratio of the Group10-12 element to aluminum in the zeolite can be about ≤1.3, such asabout ≤1.2, ≤about 1.0, or ≤about 0.8. In any embodiment, the ratio ofthe Group 10-12 element to aluminum is about 0.2 to about 1.2, about 0.3to about 1.0, or about 0.4 to about 0.8. Still further additionally oralternately, the amount of Group 10-12 element can be expressed as aweight percentage of the self-bound or unbound zeolite, such as having≤about 0.1 wt %, ≥about 0.25 wt %, ≥about 0.5 wt %, ≥about 0.75 wt %, or≥about 1.0 wt % of Group 10-12 element. Additionally or alternatively,the amount of Group 10-12 element can be present in an amount of ≤about20 wt %, such as ≤about 10 wt %, ≤about 5 wt %, ≤about 2.0 wt %, ≤about1.5 wt %, ≤about 1.2 wt %, ≤about 1.1 wt %, or ≤about 1.0 wt %. In anyembodiment, the amount of Group 10-12 element may be about 0.25 to about10 wt %, about 0.5 to about 5.0 wt %, about 0.75 to about 2.0 wt %, orabout 1.0 to about 1.5 wt %, based on the total weight of the catalystcomposition excluding the weight of any binder if present.

The catalyst compositions can optionally also include a Group 15element, e.g., phosphorous, arsenic, antimony, bismuth, and combinationsthereof, in addition to the transition metal, particularly phosphorous.

The Group 15 element can be incorporated into the catalyst compositionin any of the same manners described for incorporation of the Group10-12 element. Any source of convenient source of the Group 15 elementmay be used, e.g., phosphoric acid (H₃PO₄) or ammonium dihydrogenphosphate (NH₄H₂PO₄). Typically, the catalyst composition can have amolar ratio of Group 15 to Group 10-12 element of about 0.1 to about 10.In any embodiment, the molar ratio of Group 15 to Group 10-12 elementmay be ≤about 10, ≤about 9.0, ≤about 8.0, ≤about 7.0, ≤about 6.0, ≤about 5.0, ≤about 4.0, ≤about 3.0, ≤about 2.5, ≤about 1.0, ≤about 0.5,≤about 0.4, ≤about 0.3, ≤about 0.2, or ≤about 0.1. Additionally oralternatively, the molar ratio of Group 15 to Group 10-12 element may be≥about 0.1, ≥about 0.2, ≥about 0.3, ≥about 0.4, ≥ about 0.5, ≥about 1.5,≥about 2.0, ≥about 3.0, ≥about 4.0, ≥about 5.0, ≥about 6.0, ≥ about 7.0,≥about 8.0, ≥about 9.0, or ≥about 10. Ranges of the molar ratio of Group15 to Group 10-12 element expressly disclosed include combinations ofany of the above-enumerated upper and lower limits, e.g., about 0.2 toabout 9.0, about 0.4 to about 8.0, about 0.6 to about 6.0, about 0.8 toabout 4.0, about 1.0 to about 3.0, about 1.5 to about 2.5, etc.Additionally or alternatively, the amount of Group 15 element can bepresent in an amount of about ≤5.0 wt %, such as ≤about 2.5 wt %, ≤about1.0 wt %, ≤about 0.75 wt %, ≤about 0.50 wt %, ≤about 0.25 wt %, or≤about 0.1 wt %. In any embodiment, the amount of Group 15 element maybe about 0.1 to about 5.0 wt %, about 0.25 to about 2.0 wt %, about 0.5to about 1.0 wt %, or about 1.0 wt %, based on the total weight of thecatalyst composition excluding the weight of any binder if present.Where the zeolite is a SAPO and the Group 15 element includesphosphorous, the molar amounts and weight percentages of the phosphorousrecited in this paragraph shall exclude the amount of phosphorousattributed to the SAPO zeolite.

In one embodiment, the catalyst can be modified with up to about 3 wt %phosphorous for improved stability and up to about 3 wt % zinc forimproved aromatics yield.

Additionally or alternately, the catalyst composition can besubstantially free of phosphorous. A catalyst composition substantiallyfree of phosphorous can contain about 0.01 wt % of phosphorous or less,such as less than about 0.005 wt % or less than about 0.001 wt % ofphosphorous. A catalyst composition substantially free of phosphorouscan be substantially free of intentionally added phosphorous orsubstantially free of both intentionally added phosphorous as well asphosphorous present as an impurity in a reagent for forming the catalystcomposition. Additionally or alternately, the catalyst composition cancontain no added phosphorous, such as containing no intentionally addedphosphorous and/or containing no phosphorous impurities to within thedetection limits of standard methods for characterizing a reagent and/ora resulting zeolite.

Additionally or alternatively, the catalyst compositions may include atleast one Group 2 and/or a Group 3 element. As used herein the term“Group 3” is intended to include elements in the Lanthanide series ofthe Periodic Table. In any embodiment, one or more Group 2 elements(e.g., Be, Mg, Ca, Sr, Ba and Ra) may be used. In other embodiments, oneor more Group 3 elements (e.g., Sc and Y) can be used, including orcomprising a Lanthanide (e.g., La, Ce, Pr, Nd, Sm, Eu, Gd, Tb, Dy, Ho,Er, Tm, Yb, and Lu). While an Actinide may be used, such elements arenot believed to offer any particular advantage. When present, the totalweight of the at least one Group 2 and/or Group 3 elements can be fromabout 0.1 to about 20 wt %, based on the total weight of the catalystcomposition excluding the weight of any binder if present. In anyembodiment, the amount of the at least one Group 2 and/or a Group 3element may be about 0.25 to about 10 wt %, about 0.5 to about 5.0 wt %,about 0.75 to about 2.0 wt %, or about 1.0 to about 1.5 wt %. Thepresence of Group 2 and/or Group 3 element is believed to help reducecoke formation.

The catalyst composition can employ the zeolite in its originalcrystalline form or after formulation into catalyst particles, such asby extrusion. A process for producing zeolite extrudates in the absenceof a binder is disclosed in, for example, U.S. Pat. No. 4,582,815, theentire contents of which are incorporated herein by reference.Advantageously, the Group 15 element, the Group 10-12 element, and/orthe at least one Group 2 and/or Group 3 element can be incorporatedafter formulation of the zeolite (such as by extrusion) to formself-bound catalyst particles. Optionally, a self-bound catalyst can besteamed after extrusion.

Thus, embodiments of the catalyst compositions described herein canfurther be characterized by at least one, for example at least two, oradvantageously all, of the following properties:

(a) a mesoporosity (i.e., mesoporous surface area or surface areaexternal to the zeolite) of ≤about 20 m²/g, e.g., ≤about 30 m²/g, ≤about40 m²/g, ≤about 50 m²/g, ≤about 60 m²/g, ≤about 70 m²/g, ≤about 80 m²/g,≤about 90 m²/g, ≤about 100 m²/g, or ≤about 200 m²/g. Additionally oralternatively, the mesoporous surface area may be ≤about 500 m²/g, e.g.,≤about 400 m²/g, ≤about 300 m²/g, ≤about 200 m²/g, or ≤about 100 m²/g.Exemplary such ranges for the mesoporous surface can include about 20 to500 m²/g, about 20 to about 400 m²/g, about 20 to about 300 m²/g, about20 to about 200 m²/g, about 20 to about 100 m²/g, about 20 to about 90m²/g, about 20 to about 80 m²/g, about 20 to about 70 m²/g, about 20 toabout 60 m²/g, about 20 to about 50 m²/g, about 30 to about 200 m²/g,about 30 to about 100 m²/g, about 40 to about 100 m²/g, about 50 toabout 100 m²/g, about 60 to about 100 m²/g, about 70 to about 100 m²/g,etc.;

(b) a microporous surface area of ≥about 100 m²/g, e.g., ≥about 200m²/g, ≥about 300 m²/g, ≥about 340 m²/g, ≥about 350 m²/g, ≥about 360m²/g, or ≥about 370 m²/g. Additionally or alternatively, the microporoussurface area may be ≤about 1000 m²/g, e.g., ≤about 750 m²/g, ≤about 600m²/g, or ≤about 500 m²/g. Exemplary such ranges can include about 100 toabout 1000 m²/g, about 200 to about 1000 m²/g, about 300 to about 1000m²/g, about 340 to about 1000 m²/g, about 350 to about 1000 m²/g, about360 to about 1000 m²/g, about 370 to about 1000 m²/g, about 100 to about750 m²/g, about 200 to about 750 m²/g, about 300 to about 750 m²/g,about 340 to about 750 m²/g, about 350 to about 750 m²/g, about 360 toabout 750 m²/g, about 370 to about 750 m²/g, about 360 to about 600m²/g, or about 350 to about 500 m²/g, etc.; and/or

(c) a diffusivity for 2,2-dimethylbutane of ≥about 1.0×10⁻² sec⁻¹, e.g.,≥about 1.10×10⁻² sec⁻¹, ≥about 1.15×10⁻² sec⁻¹, ≥about 1.20×10⁻² sec⁻¹,≥about 1.25×10⁻² sec⁻¹, or ≥about 1.50×10⁻² sec⁻¹ Additionally oralternatively, the diffusivity for 2,2-dimethylbutane may be ≤about3.00×10⁻² sec⁻¹, ≤about 2.75×10⁻² sec⁻¹, ≤about 2.50×10⁻² sec⁻¹ or≤about 2.00×10⁻² sec⁻¹. Exemplary such ranges can include about 1.0×10⁻²sec⁻¹ to about 3.00×10⁻² sec⁻¹, about 1.25×10⁻² to about 3.00×10⁻²sec⁻¹, about 1.50×10⁻² to about 2.00×10⁻² sec⁻¹, etc., when measured ata temperature of about 120° C. and a 2,2-dimethylbutane pressure ofabout 60 torr (about 8 kPa).

Of these properties, mesoporosity and diffusivity for 2,2-dimethylbutaneare determined by a number of factors for a given zeolite, including thecrystal size of the zeolite. Microporous surface area is determined bythe pore size of the zeolite and the availability of the zeolite poresat the surfaces of the catalyst particles. Producing a zeolite catalystwith the desired low (minimum) mesoporosity, microporous surface area,and 2,2-dimethylbutane diffusivity would be well within the expertise ofanyone of ordinary skill in zeolite chemistry. It is noted that mesoporesurface area and micropore surface area can be characterized, forexample, using adsorption-desorption isotherm techniques within theexpertise of one of skill in the art, such as the BET (Brunauer EmmetTeller) method.

It is noted that the micropore surface area can be characterized forzeolite crystals or a catalyst formed from the zeolite crystals. Invarious aspects, the micropore surface area of a self-bound catalyst ora catalyst formulated with a separate binder can be ≥about 100 m²/g,e.g., ≥about 200 m²/g, ≥about 300 m²/g, ≥about 340 m²/g, ≥about 350m²/g, ≥about 360 m²/g, or ≥about 370 m²/g. Additionally oralternatively, the microporous surface area may be ≤about 1000 m²/g,e.g., ≤about 750 m²/g, ≤about 600 m²/g, or ≤about 500 m²/g. Exemplarysuch ranges can include about 100 to about 1000 m²/g, about 200 to about1000 m²/g, about 300 to about 1000 m²/g, about 340 to about 1000 m²/g,about 350 to about 1000 m²/g, about 360 to about 1000 m²/g, about 370 toabout 1000 m²/g, about 100 to about 750 m²/g, about 200 to about 750m²/g, about 300 to about 750 m²/g, about 340 to about 750 m²/g, about350 to about 750 m²/g, about 360 to about 750 m²/g, about 370 to about750 m²/g, about 360 to about 600 m²/g, or about 350 to about 500 m²/g,etc. Typically, a formulation of zeolite crystals into catalystparticles (either self-bound or with a separate binder) can result insome loss of micropore surface area relative to the micropore surfacearea of the zeolite crystals. Thus, in order to provide a catalysthaving the desired micropore surface area, the zeolite crystals can alsohave a micropore surface area of ≥about 100 m²/g, e.g., ≥about 200 m²/g,≥about 300 m²/g, ≥about 340 m²/g, ≥about 350 m²/g, ≥about 360 m²/g, or≥about 370 m²/g. Additionally or alternatively, the microporous surfacearea may be ≤about 1000 m²/g, e.g., ≤about 750 m²/g, ≤about 600 m²/g, or≤about 500 m²/g. Exemplary such ranges can include about 100 to about1000 m²/g, about 200 to about 1000 m²/g, about 300 to about 1000 m²/g,about 340 to about 1000 m²/g, about 350 to about 1000 m²/g, about 360 toabout 1000 m²/g, about 370 to about 1000 m²/g, about 100 to about 750m²/g, about 200 to about 750 m²/g, about 300 to about 750 m²/g, about340 to about 750 m²/g, about 350 to about 750 m²/g, about 360 to about750 m²/g, about 370 to about 750 m²/g, about 360 to about 600 m²/g, orabout 350 to about 500 m²/g, etc. As a practical matter, the microporesurface area of a zeolite crystal and/or a corresponding self-bound orbound catalyst as described herein can be ≤about 1000 m²/g, andtypically ≤about 750 m²/g. Additionally or alternately, the microporesurface area of a catalyst (self-bound or with a separate binder) can be≤about 105% of the micropore surface area of the zeolite crystals in thecatalyst, and typically ≤about 100% of the micropore surface area of thezeolite crystals in the catalyst, such as from about 80% to about 100%of the micropore surface area of the zeolite crystals in the catalyst.For example, the micropore surface area of a catalyst can be ≥about 80%of the micropore surface area of the zeolite crystals in the catalyst,such as ≥about 85%, ≥about 90%, ≥about 95%, ≥about 97%, or ≥about 98%,and/or ≤about 100%, ≤about 99%, ≤about 98%, ≤about 97%, or ≤about 95%.

Additionally or alternatively, the diffusivity for 2,2-dimethylbutane ofa catalyst composition (self-bound or with a separate binder) can be≤about 105% of the diffusivity for 2,2-dimethylbutane of the zeolitecrystals in the catalyst, and typically ≤about 100% or of thediffusivity for 2,2-dimethylbutane of the zeolite crystals in thecatalyst, such as from about 80% to about 100% of the diffusivity for2,2-dimethylbutane of the zeolite crystals in the catalyst. For example,the diffusivity for 2,2-dimethylbutane of a catalyst can be ≥about 80%of the diffusivity for 2,2-dimethylbutane of the zeolite crystals in thecatalyst, such as ≥about 85%, ≥about 90%, ≥about 95%, ≥about 97%, or≥about 98%, and/or ≤about 100%, ≤about 99%, ≤about 98%, ≤about 97%, or≤about 95%.

Additionally or alternatively, the catalyst composition comprisesparticles having a size ≥about 0.01 μm, ≥about 0.05 μm, ≥about 0.08 μm,≥about 0.10 μm, ≥about 0.20 μm, or ≥about 0.50 μm. Likewise, thecatalyst composition may comprise particles wherein the upper limit is≤about 0.6 μm, ≤about 0.5 μm, ≤about 0.4 μm, ≤about 0.3 μm, ≤about 0.2μm, ≤about 0.1 μm, or ≤about 0.05 μm. In any embodiment, the catalystmay comprise particles having a size of about 0.01 μm to about 0.6 μm,about 0.02 to about 0.50 μm, about 0.03 to about 0.40 μm etc. As usedherein the term “size” means either the diameter of approximatelyspherical particles or, where a particle has another shape, the averageof the longest dimension and the dimension orthogonal thereto. Particledimensions and size can be determined by any suitable means, typicallymicroscopically, using a representative number of particles. “Size” mayrefer to self-bound particles or particles including a binder, or thoseformed by extrusion of other method.

Additionally or alternatively, catalyst compositions herein may bedescribed by a particle size distribution, D_(x), ≤about 1.0 μm, ≤about0.5 μm, ≤about 0.40 μm, ≤about 0.20 μm, ≤about 0.10 μm, ≤about 0.05 μm,or ≤about 0.01 μm, where x is 50, 90, or 95. The particle sizedistribution, may also be ≥about 1.0 μm, ≥about 0.8 μm, ≥about 0.5 μm,≥about 0.20 μm, ≥about 0.10 μm, ≥about 0.05 μm, ≥about 0.01 μm. In anyembodiment, the particle size distribution, D_(x), may be about 0.01 toabout 0.60 μm, about 0.02 to about 0.50 μm, about 0.03 to about 0.40 μm,about 0.01 to about 0.05 μm, about 0.10 to about 0.60 μm, about 0.2 toabout 0.5 μm, or about 0.3 to about 0.4 μm. The particle sizedistribution, D_(x), means that at least x number percent of theparticles have a size, as defined above, in the recited range. Forexample, a catalyst composition described as having a D₉₀ of 0.10 to0.60 means that at least 90 number percent of the particles have a sizebetween 0.10 and 0.60 μm. In any embodiment, the particle size may berelatively narrow, i.e., D₉₀ or D₉₅ may be preferred, i.e., a D₉₀ or D₉₅of ≤about 1 μm, ≤about 0.5 μm, or ≤about 0.4 μm, about 0.01 to about0.60 μm, about 0.02 to about 0.50 μm, about 0.03 to about 0.40 μm, about0.01 to about 0.05 μm, about 0.10 to about 0.60 μm, about 0.2 to about0.5 μm, or about 0.30 to about 0.40 μm.

In some aspects, the catalyst composition can have an alpha value of atleast about 10, such as at least about 20 or at least about 50.Additionally or alternatively, the catalyst composition can have analpha value of ≤about 1000, ≤about 800, ≤ about 700, or ≤about 600;e.g., about 10 to about 1000, about 10 to about 800, or about 50 to 700.The alpha value of a catalyst composition is a measure of the acidactivity of a zeolite catalyst as compared with a standardsilica-alumina catalyst. The alpha test is described in U.S. Pat. No.3,354,078; in the Journal of Catalysis at vol. 4, p. 527 (1965), vol. 6,p. 278 (1966), and vol. 61, p. 395 (1980), each incorporated herein byreference as to that description. The experimental conditions of thetest used herein include a constant temperature of about 538° C. and avariable flow rate as described in detail in the Journal of Catalysis atvol. 61, p. 395. The higher alpha values correspond with a more activecracking catalyst.

Catalyst Binders

A catalyst composition as described herein can employ a transitionmetal-enhanced zeolite in its original crystalline form, or the crystalscan be formulated into catalyst particles, such as by extrusion. Oneexample of binding zeolite crystals to form catalyst particles is toform a self-bound catalyst. A process for producing zeolite extrudatesin the absence of a binder is disclosed in, for example, U.S. Pat. No.4,582,815, the entire contents of which are incorporated herein byreference.

As another example of forming a self-bound catalyst, the followingprocedure describes a representative method for forming self-bound ZSM-5catalyst particles. It is noted that the absolute values in gramsprovided below should be considered as representative of using anappropriate ratio of the various components. ZSM-5 crystal (such asabout 1,400 grams on a solids basis) can be added to a mixer and drymulled. Then, (approximately 190 grams of deionized) water can be addedduring mulling. After about 10 minutes, (about 28 grams of about 50 wt%) caustic (solution) mixed with (about 450 grams of deionized) watercan be added to the mixture and mulled for an additional about 5minutes. The mixture can then be extruded into (˜ 1/10″) quadralobes.The extrudates can be dried overnight (for about 8-16 hours at about250° F. (about 121° C.)) and then calcined in nitrogen (for about 3hours at about 1000° F. (about 538° C.)). The extrudates can then beexchanged twice with (an ˜1N solution of) ammonium nitrate. Theexchanged crystal can be dried overnight (for about 8-16 hours at about250° F. (about 121° C.)) and then calcined in air (for about 3 hours atabout 1000° F. (about 538° C.)). This can result in self-bound catalyst.Based on the exchange with ammonium nitrate and subsequent calcinationsin air, the ZSM-5 crystals in such a self-bound catalyst can correspondto ZSM-5 with primarily hydrogen atoms at the ion exchange sites in thezeolite. Thus, such a self-bound catalyst is sometimes described asbeing a self-bound catalyst that can include H-ZSM-5.

To form a transition metal-enhanced catalyst, a self-bound catalyst asdescribed above can be impregnated via incipient wetness with a solutioncontaining the desired metal for impregnation, such as Zn and/or Cd.(Other methods for incorporating a transition metal into the catalyst,such as ion exchange, can be used in place of or in addition to such animpregnation.) The impregnated crystal can then be dried overnight (forabout 8-16 hours at about 250° F. (about 121° C.)), followed bycalcination in air (for about 3 hours at about 1000° F. (about 538°C.)). More generally, a transition metal can be incorporated into theZSM-5 crystals and/or catalyst at any convenient time, such as before orafter ion exchange to form H-ZSM-5 crystals, or before or afterformation of a self-bound extrudate.

As an alternative to forming self-bound catalysts, zeolite crystals canbe combined with a binder to form bound catalyst compositions containinga relatively small amount of binder. Suitable binders for zeolite-basedcatalysts can include various inorganic oxides, such as silica, alumina,zirconia, titania, silica-alumina, cerium oxide, magnesium oxide, orcombinations thereof. Generally, a binder can be present in an amount of0 to about 80 wt %, ≤about 65 wt %, ≤about 40 wt %, ≤about 35 wt %,≤about 25 wt %, or ≤20 wt %, based on the total weight of the catalystcomposition. Additionally or alternatively, the binder may in anyembodiment be present in an amount of 0 wt. %, ≥about 1.0 wt %, ≥about5.0 wt %, ≥about 10 wt %, or ≥about 15 wt %, e.g., 0 to about 80 wt %,about 5.0 to about 40 wt %, about 10 to about 35, about 10 to about 25,or about 15 to about 20 wt %. In any embodiment, only a relatively smallamount of binder may be present, e.g., an upper limit of about 5.0 wt %,about 2.5 wt %, or about 1.0 wt % and a lower limit of about 0.1 wt %,about 0.5 wt %, about 1.0 wt %, such as 0.1 to 5.0 wt %, 0.5 to 2.5 wt%, 0.5 to 1.0 wt %, or 0.1 to 1.0 wt %. Combining the zeolite and thebinder can generally be achieved, for example, by mulling an aqueousmixture of the zeolite and binder and then extruding the mixture intocatalyst pellets. A process for producing zeolite extrudates using asilica binder is disclosed in, for example, U.S. Pat. No. 4,582,815.Optionally, a bound catalyst can be steamed after extrusion.

In some aspects, a binder can be used that is substantially free ofalumina, such as a binder that is essentially free of alumina. In thisdescription, a binder that is substantially free of alumina is definedas a binder than contains ≤about 10 wt % alumina, such as ≤about 7.0 wt%, ≤about 5.0 wt %, or ≤about 3.0 wt %. A binder that is essentiallyfree of alumina is defined as a binder that contains ≤about 1.0 wt %,such as about ≤0.5 wt %, or ≤about 0.1 wt %. Additionally oralternately, a binder can be used that contains no intentionally addedalumina and/or that contains no alumina within conventional detectionlimits for determining the composition of the binder and/or the reagentsfor forming the binder. Although alumina is commonly used as a binderfor zeolite catalysts, due in part to ease of formulation ofalumina-bound catalysts, in some aspects the presence of alumina in thebinder can reduce and/or inhibit the activity of a catalyst compositionfor converting methanol to aromatics. For example, for a catalyst wherethe Group 10-12 and/or Group 15 is incorporated into the catalyst afterformulation of the bound catalyst (such as by extrusion), the Group10-12 and/or Group 15 element may have an affinity for exposed aluminasurfaces relative to exposed zeolite surfaces, leading to increasedinitial deposition and/or migration of such elements to regions of thebound catalyst with an alumina surface in favor of regions with azeolite surface. Additionally or alternately, alumina-bound catalystscan tend to have low micropore surface area, meaning that the amount ofavailable zeolite surface available for receiving a Group 10-12 elementand/or Group 15 element may be undesirably low.

In some aspects, a binder for formulating a catalyst can be selected sothat the resulting bound catalyst has a micropore surface area of atleast about 3400 m²/g, such as at least about 350 m²/g or at least about370 m²/g or at least about 290 m²/g, such as at least about 300 m²/g orat least about 310 m²/g. Examples of a suitable binder for forming boundcatalysts with a desirable micropore surface area is an alumina orsilica binder. Optionally but preferably, a suitable binder can be abinder with a surface area of about 200 m²/g or less, such as about 175m²/g or less or about 150 m²/g or less. Without being bound by anyparticular theory, it is believed that catalysts formed using highsurface area binders (such as high surface area alumina binders) canhave an increased tendency for deposited added element(s) to migrate tothe binder, rather than remaining associated with the zeolite. Unlessotherwise specified, the surface area of the binder is defined herein asthe combined micropore surface area and mesopore surface area of thebinder.

As an example of forming a bound catalyst, the following proceduredescribes a representative method for forming alumina bound ZSM-5catalyst particles. ZSM-5 crystal and an alumina binder, such as analumina binder having a surface area of about 200 m²/g or less, can beadded to a mixer and mulled. Additional deionized water can be addedduring mulling to achieve a desired solids content for extrusion.Optionally, a caustic solution can also be added to the mixture andmulled. The mixture can then be extruded into a desired shape, such as ˜1/10″ quadralobes. The extrudates can be dried overnight (for about 8-16hours at about 250° F. (about 121° C.)) and then calcined in nitrogen(for about 3 hours at about 1000° F. (about 538° C.)). The extrudatescan then be exchanged twice with (an ˜1N solution of) ammonium nitrate.The exchanged crystal can be dried overnight (for about 8-16 hours atabout 250° F. (about 121° C.)) and then calcined in air (for about 3hours at about 1000° F. (about 538° C.)). This can result in an aluminabound catalyst. Based on the exchange with ammonium nitrate andsubsequent calcinations in air, the ZSM-5 crystals in such a boundcatalyst can correspond to ZSM-5 with primarily hydrogen atoms at theion exchange sites in the zeolite. Thus, such a bound catalyst issometimes described as being a bound catalyst that can include H-ZSM-5.

To form a transition metal-enhanced catalyst, a bound catalyst can beimpregnated via incipient wetness with a solution containing the desiredmetal for impregnation, such as Zn and/or Cd. The impregnated crystalcan then be dried overnight (for about 8-16 hours at about 250° F.(about 121° C.)), followed by calcination in air (for about 3 hours atabout 1000° F. (about 538° C.)). More generally, a transition metal canbe incorporated into the ZSM-5 crystals and/or catalyst at anyconvenient time, such as before or after ion exchange to form H-ZSM-5crystals, or before or after formation of a bound extrudate. In someaspects that can be preferred from a standpoint of facilitatingmanufacture of a bound zeolite catalyst, the transition metal can beincorporated into the bound catalyst (such as by impregnation or ionexchange) after formation of the bound catalyst by extrusion or anotherconvenient method.

The invention can additionally or alternately include one or more of thefollowing embodiments.

Embodiment 1

A process for converting an oxygenate feedstock to a C5+ gasolineproduct comprising: feeding the oxygenate feedstock to a fluidized bedreactor under conditions to convert the oxygenate feedstock to ahydrocarbon mixture in a reactor effluent, wherein the fluid bed reactorcomprises: a catalyst; and at least one packing layer; cooling thereactor effluent comprising the hydrocarbon mixture and condensing aportion of the reactor effluent to form a mixed phase effluent;separating the mixed phase effluent into an aqueous liquid phase, ahydrocarbon gas phase and a hydrocarbon liquid phase; and separating aC⁴⁻ light gas comprising C₂-C₄ olefins and the C₅₊ gasoline product fromthe hydrocarbon gas phase and the hydrocarbon liquid phase.

Embodiment 2

The process of embodiment 1, wherein the temperature in the fluidizedbed reactor is about 600° F. to about 900° F. and/or wherein thepressure in the fluidized bed reactor is about 25 psig to about 400psig.

Embodiment 3

The process of embodiment 1 or 2, wherein the catalyst comprises azeolite such as ZSM-5.

Embodiment 4

The process of any one of the previous embodiments, wherein theoxygenate feedstock comprises methanol and/or dimethylether.

Embodiment 5

The process of any one of the previous embodiments, wherein thefluidized bed reactor comprises at least two packing layers.

Embodiment 6

The process of any one of the previous embodiments, wherein the yield ofC₅₊ gasoline product is at least 80 wt %, e.g., from about 80 wt % toabout 90 wt %.

Embodiment 7

The process of any one of the previous embodiments, further comprisingremoving catalyst fines from the reactor effluent.

Embodiment 8

The process of any one of the previous embodiments, further comprisingremoving heat from the fluidized bed reactor internally or externally.

Embodiment 9

The process of any one of the previous embodiments, further comprisingrecycling the C⁴⁻ light gas comprising C₂-C₄ olefins, optionally to thefluidized bed reactor, to a second reactor, and/or to a catalyst cooler,under conditions to convert the C₂-C₄ olefins to C₅₊ gasoline product.

Embodiment 10

The process of embodiment 9, wherein the recycling ratio is about 3.

Embodiment 11

The process of any one of the previous embodiments, further comprisingsending the C₅₊ gasoline product to storage.

Embodiment 12

The process of any one of the previous embodiments, further comprisingregenerating the catalyst.

Embodiment 13

An apparatus for producing a C5+ gasoline product comprising: afluidized bed reactor comprising: a fluid inlet for a feedstock; acatalyst; and at least one packing layer; a cooler for cooling thereactor effluent and condensing a portion of the reactor effluent toform a mixed phase effluent; a separator for separating the mixed phaseeffluent into a gas hydrocarbon stream, a water stream, and a liquidhydrocarbon stream; a means for transporting the reactor effluent fromthe fluid bed reactor to the separator; at least one fractionatingcolumn for producing the C5+ gasoline product; and a means fortransporting the liquid hydrocarbon stream and gas hydrocarbon stream tothe at least one fractionating column.

Embodiment 14

A process for converting an oxygenate feedstock to a C₅₊ gasolineproduct comprising: (a) feeding the oxygenate feedstock to a fluidizedbed reactor under conditions to convert the oxygenate feedstock to ahydrocarbon mixture comprising C₅₊ gasoline product in a reactoreffluent, wherein the fluidized bed reactor comprises: (i) a catalyst;and (ii) two packing layers, which separate the fluidized bed reactorinto two stages; (b) cooling the fluidized bed reactor either internallyor externally; (c) transferring spent catalyst comprising coke to an airstream in fluid connection with the fluidized bed reactor and aregenerator; (d) feeding the air stream containing spent catalyst to theregenerator and burning the coke off of the catalyst to form regeneratedcatalyst; and (e) transferring the regenerated catalyst from theregenerator to the fluidized bed reactor, wherein the regenerator is influid connection with the fluidized bed reactor.

Embodiment 15

The process of embodiment 14, wherein the fluidized bed reactor iscooled internally with a heat exchanger in each stage, and/or whereinthe fluidized bed reactor is cooled externally with a catalyst cooler influid connection with the fluidized bed reactor.

Embodiment 16

The process of embodiment 14 or 15, further comprising feeding a lightgas recycle stream into the fluidized bed reactor or combining the lightgas recycle stream with the oxygenate feedstock.

Embodiment 17

A process for converting an oxygenate feedstock to a C₅₊ gasolineproduct comprising: (a) heating the oxygenate feedstock; (b) feeding theoxygenate feedstock to a fluidized bed reactor under conditions toconvert the oxygenate feedstock to a hydrocarbon mixture comprising C₅₊gasoline product in a reactor effluent, wherein the fluidized bedreactor comprises: (i) a catalyst; and (ii) two packing layers, whichseparate the fluidized bed reactor into two stages; (c) cooling thefluidized bed reactor either internally or externally; (d) transferringthe reactor effluent to a set of two stage cyclones in fluid connectionwith the fluidized bed reactor; (e) separating reactor vapor from thecatalyst in the two stage cyclones and removing catalyst fines to afines collection unit; (f) transferring the reactor effluent to a heatexchanger in fluid connection with the fines collection unit and coolingthe reactor effluent and condensing a portion of the reactor effluentagainst incoming oxygenate feed to form a mixed phase effluent; (g)transferring the mixed phase effluent to a separator in fluid connectionwith the heat exchanger and separating the mixed phase effluent into anaqueous liquid phase, a hydrocarbon gas phase and a hydrocarbon liquidphase; (h) transferring the hydrocarbon gas phase and the hydrocarbonliquid phase to a stabilizer/de-butanizer in fluid connection with theseparator, wherein a portion C⁴⁻ light gas comprising C₂-C₄ olefins andLPG and the C₅₊ gasoline product are separated; (j) recycling a portionof the C⁴⁻ light gas; (k) transferring spent catalyst comprising coke toan air stream in fluid connection with the fluidized bed reactor and aregenerator; (m) feeding the air stream containing spent catalyst to theregenerator and burning the coke off of the catalyst to form regeneratedcatalyst; and (n) transferring the regenerated catalyst from theregenerator to the fluidized bed reactor, wherein the regenerator is influid connection with the fluidized bed reactor.

Embodiment 18

The process of embodiment 17, further comprising cooling the mixed phaseeffluent before transferring the mixed phase effluent to the separator.

Embodiment 19

The process of embodiment 17 or 18, further comprising recycling theC₄-light gas in a recycle stream to the fluidized bed reactor underconditions to convert C₂-C₄ olefins to the C₅₊ gasoline product, whereinthe recycle stream is in fluid connection with the stabilizer and thefluidized bed reactor.

Embodiment 20

The process of any one of embodiments 17-19, further comprising: (o)recycling the C⁴⁻ light gas to a second reactor and converting C₂-C₄olefins to a second hydrocarbon mixture comprising C₅₊ gasoline productin a second reactor effluent; (p) transferring the second reactoreffluent to a second set of two stage cyclones in fluid connection withthe second reactor; (q) separating reactor vapor from the catalyst inthe two stage cyclones and removing catalyst fines to a fines collectionunit; (r) transferring the second reactor effluent to a second cooler influid connection with the second fines collection unit and cooling thesecond reactor effluent and condensing a portion of the second reactoreffluent to form a second mixed phase effluent; (s) transferring thesecond mixed phase effluent to a second separator in fluid connectionwith the second cooler and separating the second mixed phase effluentinto a second aqueous liquid phase, a second hydrocarbon gas phase and asecond hydrocarbon liquid phase; (t) mixing the second hydrocarbonliquid phase with the C₅₊ gasoline product from the stabilizer to form acombined mixture and transferring the second hydrocarbon gas phase andthe combined mixture to a de-ethanizer, wherein a portion C²⁻ light gasis separated from C₃₊ product, wherein the de-ethanizer is in fluidconnection with the stabilizer and the second separator; and (u)transferring the C₃₊ product to a de-butanizer in fluid connection withthe de-ethanizer, wherein the LPG and the C₅₊ gasoline product areseparated.

Embodiment 21

A process for converting an oxygenate feedstock to a C₅₊ gasolineproduct comprising: (a) heating the oxygenate feedstock; (b) feeding theoxygenate feedstock to a fluidized bed reactor under conditions toconvert the oxygenate feedstock to a hydrocarbon mixture comprising C₅₊gasoline product in a reactor effluent, wherein the fluidized bedreactor comprises: (i) a catalyst; and (ii) two packing layers, whichseparate the fluidized bed reactor into two stages; (c) cooling thefluidized bed reactor either internally or externally; (d) transferringthe reactor effluent to a set of two stage cyclones in fluid connectionwith the fluidized bed reactor; (e) separating reactor vapor from thecatalyst in the two stage cyclones and removing catalyst fines to afines collection unit; (f) transferring the reactor effluent to a heatexchanger in fluid connection with the fines collection unit and coolingthe reactor effluent and condensing a portion of the reactor effluentagainst incoming oxygenate feed to form a mixed phase effluent; (g)transferring the mixed phase effluent to a separator in fluid connectionwith the heat exchanger and separating the mixed phase effluent into anaqueous liquid phase, a hydrocarbon gas phase and a hydrocarbon liquidphase; (h) transferring the hydrocarbon gas phase and the hydrocarbonliquid phase to a de-ethanizer in fluid connection with the separator,wherein C²⁻ light gas is separated from C₃₊ product; (j) transferringthe C₃₊ product to a de-butanizer in fluid connection with thede-ethanizer, wherein C₃/C₄ gases are separated from the C₅₊ gasolineproduct; and (k) transferring the C₃/C₄ gases to an alkylation unit influid connection with the de-butanizer, wherein olefins are converted tothe C₅₊ gasoline product.

Embodiment 22

A process for converting an oxygenate feedstock to a C₅₊ gasolineproduct comprising: (a) heating the oxygenate feedstock; (b) feeding theoxygenate feedstock to a fluidized bed reactor under conditions toconvert the oxygenate feedstock to a hydrocarbon mixture comprising C₅₊gasoline product in a reactor effluent, wherein the fluid bed reactorcomprises: (i) a catalyst; and (ii) two packing layers, which separatesthe fluidized bed reactor into two stages; (c) cooling the fluidized bedreactor either internally or externally; (d) transferring the reactoreffluent to a set of two stage cyclones in fluid connection with thefluidized bed reactor; (e) separating reactor vapor from the catalyst inthe two stage cyclones; (f) transferring the reactor effluent to a finescollection unit in fluid connection with the two stage cyclones andremoving catalyst fines; (g) transferring the reactor effluent to a heatexchanger in fluid connection with the fines collection unit and coolingthe reactor effluent and condensing a portion of the reactor effluentagainst incoming oxygenate feed to form a mixed phase effluent; (h)transferring the mixed phase effluent to a separator in fluid connectionwith the heat exchanger and separating the mixed phase effluent into anaqueous liquid phase, a hydrocarbon gas phase and a hydrocarbon liquidphase; (j) transferring the hydrocarbon gas phase and the hydrocarbonliquid phase to a dividing wall column in fluid connection with theseparator, wherein seven streams for a light gas, C₂, propylene,propane, butenes, butanes and the C₅₊ gasoline product are divided; (k)combining the streams for C₂, propylene, and butenes to form a recyclestream and, wherein the recycle streams is in fluid connection with thedividing wall column and the fluidized bed reactor; (m) feeding therecycle stream to the fluidized bed reactor under conditions to convertC₂-C₄ olefins to the C₅₊ gasoline product; and (n) combining the streamsfor propane and butanes to form LPG.

Embodiment 23

The process of embodiment 22, wherein the hydrocarbon gas phase and thehydrocarbon liquid phase are transferred by a pump to the dividing wallcolumn.

EXAMPLES

The following examples are merely illustrative, and do not limit thisdisclosure in any way.

Example 1—Baffle Testing

In a cold flow study in a 2 foot diameter fluid bed, staging baffleswere tested and evaluated by using tracer techniques. It was found thatthe staging baffles greatly reduced the gas phase back-mixing and gasby-pass. Based on the results, it is believed that the C₅₊ gasolineyield of the fluid bed MTG process can be improved by ˜2-4 wt %.

Example 2—MTG Process

An MTG fluidized bed process according to above embodiments is performedwith a catalyst consisting of 40 wt % ZSM-5 (55:1 silica:alumina ratio),25 wt % silica, 5 wt % alumina, and 30 wt % clay. The catalystproperties are shown in the Table 2 below.

TABLE 2 State Calcined LOI, % 550 C 2.3 Alpha, JGC 104 Alpha, G102 NASilica wt % 79.6 Alumina wt % 19.0 Sodium, wt % 0.12 Carbon, wt %Surface Area, m2/g 214 Re2O3, wt % 0 Fe2O3, wt % 0.32 P2O5, wt % .1Attrition Rate, Initial 7.5 %/5 hrs Attrition Rate, %/15 hrs 10.5 CBD,g/cc 0.93 Particle Size Distribution 0-20 micron 3 0-40 micron 15 0-60micron 0-80 micron 0-101 micron 0-150 microns APS, microns 85

The process conditions and feed of the process can be as follows:

-   -   (i) Pressure in the range of about 25 to about 400 psig;    -   (ii) Temperature in the range of about 600 to about 900° F.;    -   (iii) Methanol WHSV, kg/kg-hr, of about 0.2 to about 3.0; and    -   (iv) Feed can be methanol (up to about 15 wt % water), DME (up        to about 40 wt % water), or a combination thereof.

The C₅₊ yield for the MTG fluidized bed process can be shown below inTable 3 as compared to the fixed bed process.

TABLE 3 MTG Yield Fixed Bed Fluid Bed Water Yield, wt % of MeOH 56 56Hydrocarbon Product, wt % of HC Light Gas 1.9 2.7 Ethylene 0.04 5.4Propane 3.1 3.5 Propylene 0.2 5.4 i-Butane 7.5 8.5 n-Butane 1.7 1.5Butenes 0.9 5.8 C5+ Gasoline 84.6 67.2 C5+ Gasoline (including Alkylate)85.2 91.2 Gasoline Octane Numbers Research Method (R + O) 93 95 MotorMethod (M = O) 83 85

What is claimed is:
 1. A process for converting an oxygenate feedstockto a C5+ gasoline product comprising: a. heating the oxygenatefeedstock; b. feeding the oxygenate feedstock to a fluidized bed reactorunder conditions to convert the oxygenate feedstock to a hydrocarbonmixture comprising C5+ gasoline product in a reactor effluent, whereinthe fluidized bed reactor comprises: i. a catalyst; and ii. at least twopacking layers, which separate the fluidized bed reactor into stages; c.cooling the fluidized bed reactor either internally or externally; d.transferring the reactor effluent to a set of two stage cyclones influid connection with the fluidized bed reactor; e. separating reactorvapor from the catalyst in the two stage cyclones and removing catalystfines to a fines collection unit; f. transferring the reactor effluentto a heat exchanger in fluid connection with the fines collection unitand cooling the reactor effluent and condensing a portion of the reactoreffluent against incoming oxygenate feed to form a mixed phase effluent;g. transferring the mixed phase effluent to a separator in fluidconnection with the heat exchanger and separating the mixed phaseeffluent into an aqueous liquid phase, a hydrocarbon gas phase and ahydrocarbon liquid phase; h. transferring the hydrocarbon gas phase andthe hydrocarbon liquid phase to a stabilizer/de-butanizer in fluidconnection with the separator, wherein a portion C4− light gascomprising C2-C4 olefins and LPG and the C5+ gasoline product areseparated; i. recycling a portion of the C4− light gas; j. transferringspent catalyst comprising coke to an air stream in fluid connection withthe fluidized bed reactor and a regenerator; k. feeding the air streamcontaining spent catalyst to the regenerator and burning the coke off ofthe catalyst to form regenerated catalyst; and l. transferring theregenerated catalyst from the regenerator to the fluidized bed reactor,wherein the regenerator is in fluid connection with the fluidized bedreactor; m. recycling the C4− light gas to a second reactor andconverting C2-C4 olefins to a second hydrocarbon mixture comprising C5+gasoline product in a second reactor effluent; n. transferring thesecond reactor effluent to a second set of two stage cyclones in fluidconnection with the second reactor; o. separating reactor vapor from thecatalyst in the two stage cyclones and removing catalyst fines to afines collection unit; p. transferring the second reactor effluent to asecond cooler in fluid connection with the second fines collection unitand cooling the second reactor effluent and condensing a portion of thesecond reactor effluent to form a second mixed phase effluent; q.transferring the second mixed phase effluent to a second separator influid connection with the second cooler and separating the second mixedphase effluent into a second aqueous liquid phase, a second hydrocarbongas phase and a second hydrocarbon liquid phase; r. mixing the secondhydrocarbon liquid phase with the C5+ gasoline product from thestabilizer to form a combined mixture and transferring the secondhydrocarbon gas phase and the combined mixture to a de-ethanizer,wherein a portion C2− light gas is separated from C3+ product, whereinthe de-ethanizer is in fluid connection with the stabilizer and thesecond separator; and s. transferring the C3+ product to a de-butanizerin fluid connection with the de-ethanizer, wherein the LPG and the C5+gasoline product are separated.
 2. The process of claim 1, furthercomprising cooling the mixed phase effluent before transferring themixed phase effluent to the separator.
 3. The process of claim 1,further comprising recycling the C4− light gas in a recycle stream tothe fluidized bed reactor under conditions to convert C2-C4 olefins tothe C5+ gasoline product, wherein the recycle stream is in fluidconnection with the stabilizer and the fluidized bed reactor.
 4. Theprocess of claim 1, wherein the fluidized bed reactor is operated at apressure of from about 25 psig to about 400 psig.
 5. The process ofclaim 1, wherein the fluidized bed reactor is operated at a temperatureof from about 500° F. to about 900° F.
 6. A process for converting anoxygenate feedstock to a C5+ gasoline product comprising: a. heating theoxygenate feedstock; b. feeding the oxygenate feedstock to a fluidizedbed reactor under conditions to convert the oxygenate feedstock to ahydrocarbon mixture comprising C5+ gasoline product in a reactoreffluent, wherein the fluid bed reactor comprises: i. a catalyst; andii. at least two packing layers, which separates the fluidized bedreactor into stages; c. cooling the fluidized bed reactor eitherinternally or externally; d. transferring the reactor effluent to a setof two stage cyclones in fluid connection with the fluidized bedreactor; e. separating reactor vapor from the catalyst in the two stagecyclones; f. transferring the reactor effluent to a fines collectionunit in fluid connection with the two stage cyclones and removingcatalyst fines; g. transferring the reactor effluent to a heat exchangerin fluid connection with the fines collection unit and cooling thereactor effluent and condensing a portion of the reactor effluentagainst incoming oxygenate feed to form a mixed phase effluent; h.transferring the mixed phase effluent to a separator in fluid connectionwith the heat exchanger and separating the mixed phase effluent into anaqueous liquid phase, a hydrocarbon gas phase and a hydrocarbon liquidphase; i. transferring the hydrocarbon gas phase and the hydrocarbonliquid phase to a dividing wall column in fluid connection with theseparator, wherein seven streams for a light gas, C2, propylene,propane, butenes, butanes and the C5+ gasoline product are divided; j.combining the streams for C2, propylene, and butenes to form a recyclestream and, wherein the recycle streams is in fluid connection with thedividing wall column and the fluidized bed reactor; k. feeding therecycle stream to the fluidized bed reactor under conditions to convertC2-C4 olefins to the C5+ gasoline product; and l. combining the streamsfor propane and butanes to form LPG.
 7. The process of claim 6, whereinthe hydrocarbon gas phase and the hydrocarbon liquid phase aretransferred by a pump to the dividing wall column.
 8. The process ofclaim 6, wherein the fluidized bed reactor is operated at a pressure offrom about 25 psig to about 400 psig.
 9. The process of claim 6, whereinthe fluidized bed reactor is operated at a temperature of from about500° F. to about 900° F.
 10. A process for converting an oxygenatefeedstock to a C5+ gasoline product comprising: a. heating the oxygenatefeedstock; b. feeding the oxygenate feedstock to a fluidized bed reactorunder conditions to convert the oxygenate feedstock to a hydrocarbonmixture comprising C5+ gasoline product in a reactor effluent, whereinthe fluid bed reactor comprises: i. a catalyst; and ii. at least onepacking layer, which separates the fluidized bed reactor into stages; c.cooling the fluidized bed reactor either internally or externally; d.transferring the reactor effluent to a set of two stage cyclones influid connection with the fluidized bed reactor; e. separating reactorvapor from the catalyst in the two stage cyclones; f. transferring thereactor effluent to a fines collection unit in fluid connection with thetwo stage cyclones and removing catalyst fines; g. transferring thereactor effluent to a heat exchanger in fluid connection with the finescollection unit and cooling the reactor effluent and condensing aportion of the reactor effluent against incoming oxygenate feed to forma mixed phase effluent; h. transferring the mixed phase effluent to aseparator in fluid connection with the heat exchanger and separating themixed phase effluent into an aqueous liquid phase, a hydrocarbon gasphase and a hydrocarbon liquid phase; i. transferring the hydrocarbongas phase and the hydrocarbon liquid phase to a stabilizer/de-butanizerin fluid connection with the separator, wherein a portion C4− light gascomprising C2-C4 olefins and LPG and the C5+ gasoline product areseparated; j. combining the streams for C2, propylene, and butenes toform a recycle stream and; k. feeding the recycle stream to thefluidized bed reactor under conditions to convert C2-C4 olefins to theC5+ gasoline product; and l. combining the streams for propane andbutanes to form LPG.
 11. The process of claim 10, wherein the fluidizedbed reactor is operated at a pressure of from about 25 psig to about 400psig.
 12. The process of claim 10, wherein the fluidized bed reactor isoperated at a temperature of from about 500° F. to about 900° F.